Process for the production of cyclohexanedimethanol

ABSTRACT

A process is described for the production of cyclohexanedimethanol having a trans-:cis-isomer ratio greater than about 1:1 by hydrogenation of a dialkyl (e.g. dimethyl) cyclohexanedicarboxylate having a trans-:cis-isomer ratio less than about 1:1 which comprises: 
     (a) providing a hydrogenation zone containing a charge of a granular heterogeneous ester hydrogenation catalyst; 
     (b) supplying to the hydrogenation zone a vaporous feed stream containing hydrogen and a hydrogenarable material comprising a dialkyl cyclohexanedicarboxylate at an inlet temperature which is above its dew point of the mixture; 
     (c) maintaining the hydrogenation zone under temperature and pressure conditions which are conducive to effecting hydrogenation of esters; 
     (d) passing the vaporous feed stream through the hydrogenation zone; and 
     (e) recovering from the hydrogenation zone a product stream containing cyclohexanedimethanol having a trans-:cis-isomer ratio greater than 1:1. 
     In this process it is preferred to use dimethyl 1,4-cyclohexanedicarboxylate to produce 1,4-cyclohexanedimethanol.

FIELD OF THE INVENTION

This invention relates to a process for the production ofcyclohexanedimethanol, and more particularly to a process for theproduction of a mixture containing a major amount of the trans-isomer ofcyclohexanedimethanol and a minor amount of the correspondingcis-isomer.

BACKGROUND OF THE INVENTION

1,4-cyclohexanedimethanol is used to prepare highly polymeric linearcondensation polymers by reaction with terephthalic acid and is usefulas an intermediate in the preparation of certain polyester and polyesteramides. The use of 1,4-cyclohexanedimethanol for such purposes isdisclosed in, for example, U.S. Pat. No. 2,901,466. This documentteaches that the trans-isomer of polycyclohexylenedimethyleneterephthalate has a higher melting point range (315°-320° C.) than thecorresponding cis-isomer (260°-267° C.).

One method for preparing 1,4-cyclohexanedimethanol(hexahydroterephthalyl alcohol) involves the hydrogenation of diethyl1,4-cyclohexanedicarboxylate (diethyl hexahydroterephthalate) in aslurry phase reactor in the presence of a copper chromite catalyst at apressure of 3000 psia (about 206.84 bar) and a temperature of 255° C.,as is described in Example 3 of U.S. Pat. No. 2,105,664. The yield issaid to be 77.5%.

The hydrogenation of dimethyl 1,4-cyclohexanedicarboxylate (DMCD) to1,4-cyclohexanedimethanol (CHDM) is shown below in equation (1):##STR1## The two geometrical isomers of CHDM thus produced are: ##STR2##

The resulting 1,4-cyclohexanedimethanol product is a mixture of thesetwo isomers which have different melting points. As reported on page 9of the book "Fiber Chemistry" edited by Menachem Lewis and Eli M.Pearce, published by Marcel Dekker, Inc.: "Both the alicyclic ester[i.e. dimethyl 1,4-cyclohexanedicarboxylate] and the alicyclic diol[i.e.1,4-cyclohexanedimethanol] exist in two isomeric forms, cis . . .and trans . . . , that are not interconvertible without bond rupture".The passage continues later: "Control of the [cis-:trans-] ratio isimportant [in 1,4-cyclohexanedimethanol] since many polymer and fiberproperties depend on it".

The cis-isomer of 1,4-cyclohexanedimethanol has a melting point of 43°C. and the trans has a melting point of 67° C. The higher melting pointtrans-isomer is often preferred over the cis-isomer for use as a reagentin the preparation of polyester and polyester-amides if a high meltingpoint for such materials is considered desirable. As noted above, thetrans-isomer of a typical polyester, such as trans-polycyclohexylmethylterephthalate, has a higher melting point than the cis-isomer. Hence,for example, U.S. Pat. No. 5,124,435 discloses a polyester copolymer,the 1,4-cyclohexanedimethanol content of which has a trans-isomercontent of at least 80 mole %, and which has a high heat resistance. Thepreferment of trans-1,4-cyclohexanedimethanol overcis-1,4-cyclohexanedimethanol is also discussed in U.S. Pat. No.2,917,549, in U.S. Pat. No. 4,999,090 and in GB-A-988316.

A liquid phase process for the production of 1,4-cyclohexanedimethanolby plural stage hydrogenation of dimethyl terephthalate is described inU.S. Pat. No. 3,334,149. This utilises a palladium catalyst to effecthydrogenation of dimethyl terephthalate to dimethyl1,4-cyclohexanedicarboxylate, followed by use of a copper chromitecatalyst in the liquid phase to catalyse the hydrogenation of thatdiester to 1,4-cyclohexanedimethanol. In the procedure described inExample 1 of that patent specification a residence time of about 40 to50 minutes is used in the second stage of this process.

In a liquid phase process for the production of1,4-cyclohexanedimethanol, such as is disclosed in U.S. Pat. No.3,334,149, the trans-:cis-isomer ratio of the product1,4-cyclohexanedimethanol will tend towards an equilibrium value. Thisequilibrium value has been reported variously and may lie between about2.57:1 (trans-:cis- 1,4-cyclohexanedimethanol) (as reported inGB-A-988316) and about 3:1 (as reported in U.S. Pat. No. 2,917,549).However, the starting material, dimethyl 1,4-cyclohexanedicarboxylate,is generally commercially obtainable as a mixture of cis- andtrans-isomers wherein there is a preponderance of the cis-isomer. Thusin a typical commercial grade of dimethyl 1,4-cyclohexanedicarboxylatethe trans-:cis-isomer ratio is from about 0.5:1 to about 0.6:1.

Attempts to deal with the problem of the presence of an excess of theless desirable cis-1,4-cyclohexanedimethanol isomer in any process for1,4-cyclohexanedimethanol manufacture have focused on the isomerisationof the cis-isomer of cyclohexanedimethanol to the trans-isomer thereof.

U.S. Pat. No. 2,917,549 discloses a process for isomerisingcis-1,4-cyclohexanedimethanol to trans-1,4-cyclohexanedimethanol whichcomprises heating cis-1,4-cyclohexanedimethanol at a temperature of atleast 200° C. in the presence of an alkoxide of a lower atomic weightmetal such as lithium, sodium, potassium, calcium or aluminium. However,the process of U.S. Pat. No. 2,917,549 necessarily involves a two-stageprocess wherein the initial cis-/trans- 1,4-cyclohexanedimethanolhydrogenation product is recovered from the hydrogenation zone andsubjected to temperatures in excess of 200° C. in the presence of ametal alkoxide catalyst under an atmosphere of nitrogen. The capital andoperational costs associated with a plant designed to carry out theprocess taught in U.S. Pat. No. 2,917,549 would be undesirably high.Another disadvantage of such a plant is the associated hazard relatingto the use of metal alkoxides as catalysts in the isomerisation zone.Such catalysts are required to effect the isomerisation, which isreported not to occur under typical hydrogenation conditions usinghydrogenation catalysts such as copper/chrome or Raney nickel catalysts,according to the teaching of Example 11 of U.S. Pat. No. 2,917,549.Furthermore, steps would be required to prevent product contamination bythe metal alkoxide catalyst.

U.S. Pat. No. 4,999,090 discloses a process for the isomerisation ofcis-1,4-cyclohexanedimethanol by distillation in the presence of analkali metal hydroxide or alkoxide at a temperature of between 150° C.and 200° C. and at a pressure of between 1 mm Hg and 50 mm Hg (between1.33 millibar and 66.5 millibar). This process has very similardisadvantages to those of U.S. Pat. No. 2,917,549.

GB-A-988316teaches a process for the preparation oftrans-1,4-cyclohexanedimethanol in which a mixture of cis- andtrans-isomers of dimethyl hexahydroterephthalate (i.e. dimethyl1,4-cyclohexanedicarboxylate) is hydrogenated at elevated temperatureand pressure in the presence of a Cu/Zn catalyst.Trans-1,4-dimethylolcyclohexane (i.e. trans-1,4-cyclohexanedimethanol)is separated by crystallisation from the reaction product and then theresidual product, now enriched in cis-1,4-cyclohexanedimethanol, isrecycled to the hydrogenation zone whereupon it undergoes isomerisationto a cis-/trans- 1,4-cyclohexanedimethanol mixture. The recycleprocedure may be repeated to obtain a 1,4-cyclohexanedimethanol productcontaining the trans-isomer in substantial excess. However, the processaccording to GB-A-988316is more preferably operated under conditionssuch that recycled cis-isomer enriched product is combined with freshdimethyl 1,4-cyclohexanedicarboxylate feed on re-entry to thehydrogenation zone. The effectiveness of recycling the cis-isomer to thehydrogenation zone is largely a result of the dual function of thecopper/zinc catalyst which possesses both a hydrogenating and anisomerising catalytic action. As would be expected from thermodynamicprinciples, the isomerising action is most effective when a mixturecontaining a preponderance of the cis-isomer is recycled to thehydrogenation zone. However, recycling the cis-isomer in this way isacknowledged to cause a new problem, that of the formation of unwantedby-products, such as 1-methyl-4-hydroxymethylcyclohexane, which may beformed by operating the hydrogenation reaction under too severeconditions. To minimise the formation of such by-products, thehydrogenation zone may be operated under "relatively mild conditions",according to the teaching of GB-A-988316(see, for example page 2, lines55 to 79 of GB-A-988316). However, such mild conditions reduce theachieved conversion of dimethyl 1,4-cyclohexanedicarboxylate with theresult that, for any one pass through the hydrogenation zone, asignificant quantity of dimethyl hexahydroterephthalate (dimethyl1,4-cyclohexanedicarboxylate) remains unconverted. By the term"relatively mild conditions" is meant a temperature of at least 200° C.,preferably between 240° C. and 300° C., and a pressure of 200 to 300atmospheres (202.65 bar to 303.98 bar), according to page 2, lines 26 to32 of GB-A-988316. The use of such high pressures at these elevatedtemperatures can be hazardous, besides requiring reactors with thickwalls and flanges of special alloy constructed to withstand such extremepressures. Hence it is expensive to construct a plant to operate atpressures as high as envisaged in GB-A-988316. Furthermore it ispotentially hazardous to operate a plant operating at 200 atmospheres(202.65 bar) or above, as well as being very expensive, not only interms of the capital cost of the plant but also with regard to operatingcosts. A substantial proportion of this capital cost is associated withthe rigorous safety precautions that must be taken when operating a highpressure conventional commercial scale hydrogenation plant. It is alsoexpensive to compress gaseous streams to such high pressures and tocirculate them through the plant.

Although there is a passing reference (see page 1, line 84 ofGB-A-988316) to use of "the gaseous phase", even at temperatures of 300°C. both cis- and trans- dimethyl hexahydroterephthalate would be in theliquid phase at pressures of 200 to 300 atmospheres (202.65 bar to303.98 bar) at the hydrogen:ester ratio envisaged in the Examples. Thusin each of the Examples of GB-A-988316liquid phase conditions are used.According to Example 4, which uses a feed mixture containing dimethylhexahydroterephthalate (i.e. 1,4-dimethyl cyclohexanedicarboxylate), andmethanol such as might be used in a recycling process, the isomerspresent in the diol in the hydrogenation product are stated to representan equilibrium mixture of about 72% of the trans- and about 28% of thecis-isomer, i.e. a trans-:cis-ratio of about 2.57:1.

It is known to effect hydrogenation of certain esters and diesters inthe vapour phase. For example it has been proposed to use a reducedcopper oxide/zinc oxide catalyst for effecting hydrogenation of estersin the vapour phase. In this connection attention is directed toGB-B-2116552. Also of relevance is WO-A-90/8121.

It is further known to produce diols, such as butane-1,4-diol, bycatalytic hydrogenation of esters of dicarboxylic acids, such as adimethyl or diethyl ester of maleic acid, fumaric acid, succinic acid,or a mixture of two or more thereof. Such processes are described, forexample, in GB-A-1454440, GB-A-1464263, DE-A-2719867, U.S. Pat. No.4,032,458, and U.S. Pat. No. 4,172,961.

Production of butane 1,4-diol by vapour phase hydrogenation of adiester, typically a dialkyl ester, of a C₄ dicarboxylic acid selectedfrom maleic acid, fumaric acid, succinic acid, and a mixture of two ormore thereof has been proposed. In such a process the diester isconveniently a di-C₁ to C₄ alkyl) ester, such as dimethyl or diethylmaleate, fumarate, or succinate. A further description of such a processcan be found in U.S. Pat. No. 4,584,419, EP-A-0143634, WO-A-86/03189,WO-A-86/07358, and WO-A-88/00937.

In all of the above-mentioned vapour phase processes the esters ordiesters all have a vapour pressure which is high compared to the vapourpressure of dimethyl 1,4-cyclohexanedicarboxylate and1,4-cyclohexanedimethanol.

SUMMARY OF THE INVENTION

It is accordingly an object of the present invention to provide aprocess for the production of cyclohexanedimethanol by hydrogenation ofa dialkyl cyclohexanedicarboxylate, for example dimethylcyclohexanedicarboxylate, which can be operated with substantiallyincreased safety and operating economy at relatively low pressures.Another object of the invention is to provide a process for productionof cyclohexanedimethanol by hydrogenation of dimethylcyclohexanedicarboxylate wherein the hydrogenation step yields directlya cyclohexanedimethanol product with a higher trans-:cis-isomer ratiothan is reported to be achievable by conventional hydrogenation methods.Hence it is a still further object of the invention to avoid theincreased capital and operating costs of the prior art processesmentioned above which require the use of extreme hydrogenationconditions or a separate isomerisation step. It is also an object of thepresent invention to provide a process wherein a mixture of cis- andtrans-isomers of dimethyl 1,4-cyclohexanedicarboxylate is reactedrapidly with high conversion and high selectivity to a mixture of cis-and trans-isomers of 1,4-cyclohexanedimethanol wherein the trans-isomeris present in excess of the cis-isomer.

According to the present invention there is provided a process for theproduction of cyclohexanedimethanol having a trans-:cis -isomer ratiogreater than about 1:1 by hydrogenation of a dialkylcyclohexanedicarboxylate having a trans-:cis-isomer ratio less thanabout 1:1 which comprises:

(a) providing a hydrogenation zone containing a charge of a granularheterogeneous ester hydrogenation catalyst;

(b) supplying to the hydrogenation zone a vaporous feed stream of amixture containing hydrogen and a hydrogenarable material comprising adialkyl cyclohexanedicarboxylate at a feed temperature which is abovethe dew point of the mixture;

(c) maintaining the hydrogenation zone under temperature and pressureconditions which are conducive to effecting hydrogenation of esters;

(d) passing the vaporous feed stream through the hydrogenation zone; and

(e) recovering from the hydrogenation zone a product stream containingcyclohexanedimethanol having a desired trans-:cis-isomer ratio greaterthan 1:1.

The invention is based upon the surprising discovery that, not only isthe conversion of the starting material, i.e. dialkylcyclohexanedicarboxylate, to the product, i.e. cyclohexanedimethanol,extremely rapid under the vapour phase hydrogenation conditions used,requiring only a matter of a few seconds for substantially completeconversion to occur, but also the isomerisation of cyclohexanedimethanolthat occurs in passage through the hydrogenation zone is comparablyrapid. This is a surprising finding since two separate reactions areinvolved. Thus a high trans-:cis- 1,4-cyclohexanedimethanol ratiotypically greater than about 2.0:1 up to about 3.84:1, as well asessentially complete conversion of dialkyl 1,4-cyclohexanedicarboxylateto 1,4-cyclohexanedimethanol, can be achieved by using a residence timeof less than about a minute of the reaction mixture in the hydrogenationzone, typically in the range of from about 2 to about 15 seconds. Thisresidence time is in stark contrast to the extended residence timesrecommended in the prior art, such as the 40 to 50 minutes residencetime used in Example 1 of U.S. Pat. No. 3,334,149.

The process of the invention is operated using vaporous feed conditionswith the feed stream being supplied to the hydrogenation zone inessentially liquid free vaporous form. Hence the feed stream is suppliedto the hydrogenation zone at an inlet temperature which is above the dewpoint of the mixture. The process can be operated so that vapour phaseconditions will exist throughout the hydrogenation zone. However, whenusing dimethyl 1,4-cyclohexanedicarboxylate as starting material, theproduct, 1,4-cyclohexanedimethanol, is less volatile than the startingmaterial, dimethyl 1,4-cyclohexanedicarboxylate, so that there is thepossibility of condensation of a 1,4-cyclohexanedimethanol-rich liquidoccurring on the catalyst, particularly if the temperature of the feedstream is close to its dew point. Such condensation of a1,4-cyclohexanedimethanol-rich liquid on the catalyst is not deleteriousin the process of the invention because the heat of hydrogenation of anydimethyl 1,4-cyclohexanedicarboxylate present in the1,4-cyclohexanedimethanol-rich liquid can be dissipated by the heat sinkeffect of the 1,4-cyclohexanedimethanol. However, it is essential thatthe feed stream be at a feed temperature above its dew point at theinlet end of the catalyst bed if the advantage of vapour phase feedconditions is to be realised. The use of vapour phase feed conditions inthe process of the invention has the advantage that, compared withliquid phase operation of the process, generally lower operatingpressures can be used. This generally has a significant and beneficialeffect not only on the construction costs but also on the operatingcosts of the plant.

In the hydrogenation zone the hydrogenatable material undergoesextremely rapid hydrogenation to yield cyclohexanedimethanol, asexemplified by equation (1) above. In addition, a comparably fastisomerisation reaction occurs under the hydrogenation conditions used.Hence a hydrogenatable material rich in cis-dimethylcyclohexanedicarboxylate tends to yield a cyclohexanedimethanol productthat has a higher content of the trans-isomer thereof than thetrans-content of the starting ester.

In the case of hydrogenation of dimethyl 1,4-cyclohexanedicarboxylate wehave found that the trans-:cis-isomer of the 1,4-cyclohexanedimethanolpresent in the reaction product mixture can be as high as approximately3.84:1 under vapour phase hydrogenation conditions. This ratio issignificantly higher than the ratio of between 2.57:1 and 3:1 which isreported to be achieved under liquid phase reaction conditions at lowconversions.

In a preferred process the dialkyl cyclohexanedicarboxylate is a di-(C₁to C₄ alkyl) cyclohexanedicarboxylate, e.g. a dimethyl, diethyl, di-n-or -iso-propyl, or di-n-, -iso- or -sec-butyl cyclohexanedicarboxylate,more preferably dimethyl cyclohexanedicarboxylate. In an especiallypreferred process the dimethyl cyclohexanedicarboxylate is dimethyl1,4-cyclohexanedicarboxylate.

In a particularly preferred process the dialkyl cyclohexanedicarboxylateis dimethyl 1,4-cyclohexanedicarboxylate and the product stream of step(e) comprises a 1,4-cyclohexanedimethanol product in which thetrans-:cis-isomer ratio is in the range of from about 2:1 up to about3.84:1, and is more preferably at least about 2.6:1 up to about 3.84:1.The process of the invention thus readily permits the preparation in asingle step of a 1,4-cyclohexanedimethanol product in which thetrans-:cis-isomer ratio is in excess of about 2:1 up to about 3.84:1,for example from about 2.6:1 to about 2.7:1. It is preferred that thisratio is at least about 2.6:1. The invention permits production in asingle step of a 1,4-cyclohexanedimethanol product with atrans-:cis-isomer ratio of from about 3.1:1 up to about 3.84:1, e.g.about 3.2:1 up to about 3.7:1.

The hydrogenarable material supplied to the hydrogenation zone maycomprise substantially pure cis-dimethyl cyclohexanedicarboxylate.However, it may further comprise trans-dimethyl cyclohexanedicarboxylatein addition to the cis-isomer thereof, provided that thetrans-:cis-isomer ratio is less than 1:1. When the hydrogenarablematerial supplied to the hydrogenation zone comprises trans-dimethylcyclohexanedicarboxylate, the molar ratio of trans-dimethylcyclohexanedicarboxylate to cis-dimethyl cyclohexanedicarboxylate may bein the range of from about 0.01:1 to about 1:1, preferably in the rangeof from about 0.05:1 to about 1:1. Typically the molar ratio of thetrans-isomer to the cis-isomer in the hydrogenatable material will befrom about 0.5:1 to about 0.6:1 when the dimethylcyclohexanedicarboxylate is the 1,4-isomer.

Whilst the process is used to advantage for the hydrogenation ofdimethyl 1,4-cyclohexanedicarboxylate, it will be understood by thoseskilled in the art that the process of the invention may be equally wellapplied to the hydrogenation of any or all of dimethyl1,2-cyclohexanedicarboxylate, dimethyl 1,3-cyclohexanedicarboxylate ordimethyl 1,4-cyclohexanedicarboxylate, and mixtures of two or morethereof.

In a preferred aspect of the present invention there is provided avapour phase process for the production of cyclohexanedimethanol byhydrogenation of dimethyl cyclchexanedicarboxylate which comprises:

(a) providing a hydrogenation zone containing a charge of a granularheterogeneous ester hydrogenation catalyst;

(b) supplying to the hydrogenation zone a vaporous feed streamcontaining hydrogen and dimethyl cyclohexanedicarboxylate at an inlettemperature which is at or above the dew point of the mixture, whereinthe dimethyl cyclohexanedicarboxylate is supplied to the hydrogenationzone at a rate corresponding to a liquid hourly space velocity of fromabout 0.05 to about 4.0 h⁻¹ ;

(c) maintaining the hydrogenation zone at a combination of a temperatureof from about 150° C. to 300° C. and a pressure of from about 150 psia(about 10.34 bar) to about 2000 psia (about 137.90 bar) which maintainsboth the dimethyl cyclohexanedicarboxylate reactant and thecyclohexanedimethanol product in the vapour phase; and

(d) recovering from the hydrogenation zone a vaporouscyclohexanedimethanol product wherein the trans-:cis-isomer ratio of thecyclohexanedimethanol product is in the range of from about 2:1 to about3.8:1.

The invention also provides a vapour phase process for the production of1,4-cyclohexanedimethanol by hydrogenation of dimethyl1,4-cyclohexanedicarboxylate which comprises:

(a) providing a hydrogenation zone containing a charge of a granularheterogeneous ester hydrogenation catalyst;

(b) supplying to the hydrogenation zone a vaporous feed streamcontaining hydrogen and dimethyl 1,4-cyclohexanedicarboxylate at aninlet temperature which is at or above the dew point of the mixture,wherein the dimethyl cyclohexanedicarboxylate is supplied to thehydrogenation zone at a rate corresponding to a liquid hourly spacevelocity of from about 0.2 to about 1.0 h⁻¹ ;

(c) maintaining the hydrogenation zone at a combination of a temperatureof from about 210° C. to about 260° C. and a pressure of from about 450psia (about 31.03 bar) to about 1000 psia (about 68.95 bar) whichmaintains both the dimethyl 1,4-cyclohexanedicarboxylate reactant andthe 1,4-cyclohexanedimethanol product in the vapour phase; and

(d) recovering from the hydrogenation zone a vaporous1,4-cyclohexanedimethanol product wherein the trans-:cis-isomer ratio ofthe 1,4-cyclohexanedimethanol product is in the range of from about2.6:1 to about 3.8:1.

In the prior art hydrogenation of dimethyl cyclohexanedicarboxylate hasnormally been carried out as a high pressure, liquid phase process. Itis surprising that, despite the relatively high molecular weight and lowvapour pressure of dimethyl cyclohexanedicarboxylate at temperaturesbelow those at which thermal decomposition is likely, e.g. below about300° C., it is possible to design a viable commercial hydrogenationplant for hydrogenation of dimethyl cyclohexanedicarboxylate whichutilises vapour phase feed conditions.

A surprising discovery is that the process of the invention can generate1,4-cyclohexanedimethanol product mixtures whose trans-:cis-isomer ratiois in excess of the ratio that is the normal equilibrium ratio underliquid phase conditions. Thus, although the normal trans-:cis-isomerequilibrium ratio, under liquid phase conditions, can under favourableconditions be as high as about 3:1, the present invention, whichoperates with vaporous feed conditions, enables production of1,4-cyclohexanedimethanol with a trans-:cis-isomer ratio as high as3.84:1.

It is also surprising that the process of the invention can be operatedfor extended periods of up to several months or more, at highselectivity to the desired cyclohexanedimethanol product and at highconversion of the starting dialkyl cyclohexanedicarboxylate, since theprejudice in the art is that copper-containing catalysts would only havea short operating life and low activity at low operating pressures.

In the prior art it is made clear that the hydrogenation of dimethyl1,4-cyclohexanedicarboxylate (dimethyl hexahydroterephthalate) is liableto give significant quantities of by-products. Thus GB-A-988316acknowledges the problem caused by formation of unwanted by-products,such as 1-methyl-4-hydroxymethylcyclohexane. It is surprising to findthat the process of the invention can be operated so that, despite thepresence of a very large excess of hydrogen and the use of a very largevapour pressure of hydrogen compared to the relatively low vapourpressure of dimethyl cyclohexanedicarboxylate, the reaction proceedsrapidly with a very high conversion of dimethyl cyclohexanedicarboxylateto the desired product, i.e. cyclohexanedimethanol, but yet with a veryhigh selectivity to that product and therefore with a very low yield ofby-products. Thus under favourable conditions the conversion of dimethylcyclohexanedicarboxylate to cyclohexanedimethanol can be as high as 98mole % or higher with a selectivity to cyclohexanedimethanol of greaterthan 96 mole %. In addition it surprising to find that it is possible,under suitable reaction conditions, to convert a starting ester(dimethyl 1,4-cyclohexanedicarboxylate) that has a trans-:cis-isomerratio less than 1:1 in a single step to a cyclohexanedimethanol productwith a trans-:cis-isomer ratio that is greater than 1:1 and normally inthe range of from about 2.6:1 up to as high as about 3.84:1, a valuethat is much higher than the highest equilibrium trans-:cis-isomer ratiothat is disclosed in the prior art. It is further surprising to findthat this can be accomplished at relatively low operating temperaturesand pressures.

In a commercial plant the process will normally be operated on acontinuous basis. It may be preferred to employ at least twohydrogenation zones, each containing a charge of a heterogeneous esterhydrogenation catalyst, connected in parallel. Each of these zones iscapable of independent isolation from the supply of vaporous feedstockmixture. Hence an isolated zone may be subjected to conditionssubstantially different from those prevailing in the remaining zone orzones, for example, whereunder the catalyst charge therein may bereactivated or replaced whilst the process of the invention is continuedin the remaining zone or zones. This arrangement also permits operationunder the conditions taught in WO-A-91/01961. In this case twohydrogenation reactors connected in parallel are used. In a first phaseof operation with a fresh charge of catalyst one only of the reactors isused, the other one being in standby mode with the catalyst bathed inhydrogen. After a period of operation over which the catalyst activitymay decline somewhat the second reactor is used, whilst the first one isplaced in standby condition. After a further period of operation bothreactors are used in parallel until the time comes to replace the entirecatalyst charge.

The process of the invention is normally operated at a feed temperatureof at least about 150° C. and no higher than about 350° C. The feedtemperature is preferably in the range of from about 150° C. to about300° C., most preferably from about 200° C. to about 260° C.

The feed pressure typically is in the range of from about 150 psia(about 10.34 bar) up to about 2000 psia (about 137.90 bar). However, thebenefits and advantages of the present low pressure, process utilisingvaporous feed conditions are best realised by carrying out the processusing a feed pressure of from about 450 psia (about 31.03 bar) up toabout 1000 psia (about 68.95 bar).

The process requires that the vaporous feed stream is above its dewpoint so that the dialkyl (e.g. dimethyl) cyclohexanedicarboxylate ispresent in the vapour phase at the inlet end of the or each catalystbed. This means that the composition of the vaporous feed mixture mustbe controlled so that, under the selected operating conditions, thetemperature of the mixture at the inlet end of the or each catalyst bedis always above its dew point at the operating pressure. By the term"dew point" is meant that temperature at which a mixture of gases andvapours just deposits a fog or film of liquid. This dew point liquidwill normally contain all the condensable components of the vapourphase, as well as dissolved gases, in concentrations that satisfy theusual vapour/liquid criteria. Typically the feed temperature of thevaporous feed mixture to the hydrogenation zone is from about 5° C. upto about 10° C. or more above its dew point at the operating pressure.

A convenient method of forming a vaporous mixture for use in the processof the invention is to spray liquid dialkyl (e.g. dimethyl)cyclohexanedicarboxylate or a dialkyl cyclohexanedicarboxylate solutioninto a stream of hot hydrogen-containing gas so as to form a saturatedor partially saturated vaporous mixture. Alternatively such a vapourmixture can be obtained by bubbling a hot hydrogen-containing gasthrough a body of the liquid dialkyl cyclohexanedicarboxylate or dialkylcyclohexanedicarboxylate solution. If a saturated vapour mixture isformed it should then be heated further or diluted with more hot gas soas to produce a partially saturated vaporous mixture prior to contactwith the catalyst.

In the process of the invention the hydrogen-containing gas:dialkyl(e.g. dimethyl) cyclohexanedicarboxylate molar ratio can vary withinwide limits, depending upon the temperature and pressure. Although themajor constituent of the hydrogen-containing gas is hydrogen, othergases may also be introduced, normally in minor amount, in thehydrogen-containing gas supplied to the process, such as nitrogen,argon, methane, and carbon oxides. In order to maintain the vaporousfeed stream above its dew point at the inlet end of the or each catalystbed at the operating pressure the hydrogen-containing gas:dialkyl (e.g.dimethyl) cyclohexanedicarboxylate molar ratio is desirably at leastabout 10:1 up to about 8000:1, preferably in the range of from about200:1 to about 1000:1. When using dimethyl 1,4-cyclohexanedicarboxylateas the diester starting material, it is not, however, necessary that thevaporous mixture in contact with all parts of the or each catalyst bedshould be so far above its dew point as to prevent condensation of a1,4-cyclohexanedimethanol-rich liquid. (Cyclohexanedimethanol is lessvolatile than the ester starting material, dimethyl1,4-cyclohexanedicarboxylate).

The hydrogen-containing gas used in the process may comprise freshmake-up gas or a mixture of make-up gas and recycle gas. The make-up gascan be a mixture of hydrogen, optional minor amounts of components suchas CO and CO₂, and inert gases, such as argon, nitrogen, or methane,containing at least about 70 mole % of hydrogen. Preferably the make-upgas contains at least 90 mole %, and even more preferably at least 97mole %, of hydrogen. The make-up gas can be produced in any convenientmanner, e.g. by partial oxidation or steam reforming of natural gasfollowed by the water gas shift reaction, and CO₂ absorption, followedpossibly by methanation of at least some of any residual traces ofcarbon oxides. Pressure swing absorption can be used if a high purityhydrogen make-up gas is desired. If gas recycle is utilised in theprocess then the recycle gas will normally contain minor amounts of oneor more products of the hydrogenation reaction which have not been fullycondensed in the product recovery stage downstream from thehydrogenation zone. For example, when using gas recycle, the gas recyclestream will normally contain minor amounts of an alkanol (e.g.methanol).

Although the process of the invention is operated with the feed streamin the vapour phase, it is convenient to express the feed rate of thehydrogenarable material to the hydrogenation zone as a space velocityand to express that space velocity as a liquid hourly space velocity.Hence it is convenient to express the feed rate in terms of the ratio ofthe liquid feed rate of the hydrogenarable material to the vaporisationzone to the volume of the hydrogenation catalyst. Thus the equivalentliquid hourly space velocity of the hydrogenarable material through thehydrogenation catalyst is preferably from about 0.05 to about 4.0 h⁻¹ Inother words it is preferred to feed the liquid hydrogenarable materialto the vaporisation zone at a rate which is equivalent to, per unitvolume of catalyst, from about 0.05 to about 4.0 unit volumes ofhydrogenatable material per hour (i.e. about 0.05 to about 4.0 m³ h⁻¹per m³ of catalyst). Even more preferably the liquid hourly spacevelocity is from about 0.1 h⁻¹ to about 1.0 h⁻¹.

It will be readily apparent to those skilled in the art that the rate ofpassage of the vaporous feed stream through the hydrogenation zone willdepend upon the feed rate of the hydrogenatable material to thevaporisation zone and upon the hydrogen-containing gas:hydrogenatablematerial molar ratio.

When using dimethyl 1,4-cyclohexanedicarboxylate feed there may be usedany feedstock containing a significant quantity of cis-dimethyl1,4-cyclohexanedicarboxylate and a trans-:cis- dimethyl1,4-cyclohexanedicarboxylate isomer ratio of less than about 1:1.Suitable commercially available cis-rich grades of the feed esterdimethyl 1,4-cyclohexanedicarboxylate are high purity dimethyl1,4-cyclohexanedicarboxylate, technical grade dimethyl1,4-cyclohexanedicarboxylate, and cis-dimethyl1,4-cyclohexanedicarboxylate. The preferred feedstock for the process ofthe invention is the technical grade dimethyl1,4-cyclohexanedicarboxylate, as the high purity and cis-dimethyl1,4-cyclohexanedicarboxylate requires additional purification stages toproduce these grades.

The granular catalyst used in the process of the invention may be anycatalyst capable of catalysing hydrogenation or hydrogenolysis of anester to the corresponding alcohol or mixture of alcohols. It may beformed into any suitable shape, e.g. pellets, rings or saddles.

Typical ester hydrogenation catalysts include copper-containingcatalysts and Group VIII metal-containing catalysts. Examples ofsuitable copper-containing catalysts include reduced copper-on-aluminacatalysts, reduced copper oxide/zinc oxide catalysts, with or without apromoter, reduced manganese promoted copper catalysts, and reducedcopper chromite catalysts, with or without a promoter, while suitableGroup VIII metal catalysts include platinum catalysts and palladiumcatalysts. Suitable copper oxide/zinc oxide catalyst precursors includeCuO/ZnO mixtures wherein the Cu:Zn weight ratio ranges from about 0.4:1to about 2:1. An example is the catalyst precursor bearing thedesignation DRD 92/71. Promoted copper oxide/zinc oxide precursorsinclude CuO/ZnO mixtures wherein the Cu:Zn weight ratio ranges fromabout 0.4:1 to about 2:1 which are promoted with from about 0.1% byweight up to about 15% by weight of barium, manganese or a mixture ofbarium and manganese. Such promoted CuO/ZnO mixtures include theMn-promoted CuO/ZnO precursor available under the designation DRD 92/92.Suitable copper chromite catalyst precursors include those wherein theCu:Cr weight ratio ranges from about 0.1:1 to about 4:1, preferably fromabout 0.5:1 to about 4:1 . Catalyst precursors of this type are theprecursors available under the designation DRD 89/21 or under thedesignation PG 85/1. Promoted copper chromite precursors include copperchromite precursors wherein the Cu:Cr weight ratio ranges from about0.1:1 to about 4:1, preferably from about 0.5:1 to about 4:1, which arepromoted with from about 0.1% by weight up to about 15% by weight ofbarium, manganese or a mixture of barium and manganese. Manganesepromoted copper catalyst precursors typically have a Cu:Mn weight ratioof from about 2:1 to about 10:1 and can include an alumina support, inwhich case the Cu:Al weight ratio is typically from about 2:1 to about4:1. An example is the catalyst precursor DRD 92/89.

All of the above mentioned catalysts available under the generaldesignations DRD or PG can be obtained from Davy Research andDevelopment Limited, P.O. Box 37, Bowesfield Lane, Stockton-on-Tees,Cleveland TS18 3HA, England.

Other catalysts which can be considered for use include Pd/ZnO catalystsof the type mentioned by P. S. Wehner and B. L. Gustafson in Journal ofCatalysis 136, 420-426 (1992), supported palladium/zinc catalysts of thetype disclosed in U.S. Pat. No. 4,837,368 and U.S. Pat. No. 5,185,476,and chemically mixed copper-titanium oxides of the type disclosed inU.S. Pat. No. 4,929,777.

Further catalysts of interest for use in the process of the inventioninclude the rhodium/tin catalysts reported in A. El Mansour, J. P.Candy, J. P. Bournonville, O. A. Ferrehi, and J. M Basset, Angew. Chem.101, 360 (1989).

Any recognised supporting medium may be used to provide physical supportfor the catalyst used in the process of the invention. This support canbe provided by materials such as zinc oxide, alumina, silica,alumina-silica, silicon carbide, zirconia, titania, carbon, a zeolite,or any suitable combination thereof.

The catalysts that are particularly preferred for use in the process ofthe invention are the reduced forms of the copper chromite, promotedcopper chromite, and manganese promoted copper catalyst precursorsdescribed hereinabove.

According to one preferred procedure the process of the invention iscarried out using at least two hydrogenation reactors connected inparallel, each zone containing a respective charge of the hydrogenationcatalyst. The vaporous feed stream may, in a first phase of operationwith a fresh catalyst charge, be supplied to one reactor only whilst thesecond reactor is in standby mode and then, in a second phase ofoperation, be supplied to the second reactor whilst the first reactor isin standby mode. In each of these first two phases the initially freshcatalyst charge in the respective reactor will become somewhatdeactivated. In a third phase of operation the vaporous feed mixture issupplied to both reactors simultaneously, as taught in WO-A-91/01961. Inthis way the effective life of the total catalyst charge can beextended.

It is also contemplated to use a hydrogenation zone which comprises twoor more hydrogenation reactors connected in series.

The or each hydrogenation zone may comprise a shell-and-tube reactorwhich may be operated under isothermal, or near isothermal, conditionswith the catalyst in the tubes and the coolant in the shell or viceversa. Usually, however, it will be preferred to use adiabatic reactorssince these are cheaper to construct and install than shell-and-tubereactors. Such an adiabatic reactor may contain a single charge of ahydrogenation catalyst or may contain two or more beds of catalyst, orbeds of different hydrogenation catalysts. If desired, external orinternal inter-bed heat exchangers may be provided in order to adjustthe inlet temperature to one or more beds of catalyst downstream fromthe inlet to the adiabatic hydrogenation reactor.

In an alternative procedure the plant includes at least twohydrogenation zones, each containing a charge of granular hydrogenationcatalyst, and the vaporous feed stream is supplied to at least one ofthe hydrogenation zones, in a first phase of operation, while at leastone other hydrogenation zone is supplied with a stream ofhydrogen-containing gas thereby to reactivate the charge ofhydrogenation catalyst therein. In a second phase of operation the atleast one other hydrogenation zone is supplied with the vaporous feedstream while the at least one hydrogenation zone is supplied with astream of hydrogen-containing gas thereby to reactivate the charge ofhydrogenation catalyst therein.

In this procedure the or each hydrogenation zone that is not on line issupplied with a stream of hydrogen-containing gas thereby to reactivatethe catalyst charge. Normally it will be preferred to carry out thisreactivation at elevated temperature, typically at a temperature ofabout 100° C. or more up to about 350° C. In this reactivation step theinlet temperature to the respective hydrogenation zone or zones may belower than, e.g. about 10° C. to about 50° C. lower than, substantiallyequal to, or higher than, e.g. about 10° C. to about 50° C. higher than,the feed temperature of the vaporous feed stream to the on linehydrogenation zone or zones. The stream of hydrogen-containing gas usedin such a reactivation step may comprise a hot stream of recycle andmake-up gas.

In one form of this alternative procedure there is recovered from thezone or zones undergoing reactivation of the catalyst a stream ofhydrogen-containing gas which is admixed with a vaporoushydrogen-containing stream of the dialkyl (e.g. dimethyl)cyclohexanedicarboxylate to form the vaporous feed stream to the on linezone or zones. In another form there is recovered from the zone or zoneswhose catalyst is undergoing reactivation a stream ofhydrogen-containing gas which is admixed with the reaction productstream from the on line zone or zones. In a still further form of thisalternative procedure there is recovered from the zone or zonesundergoing catalyst reactivation a stream of hydrogen-containing gaswhich is used to vaporise the dialkyl (e.g. dimethyl)cyclohexanedicarboxylate to form a vaporous hydrogen-containing streamof the dialkyl cyclohexanedicarboxylate. It is further convenient toform the vaporous feed stream to the on line zone or zones by admixinghot recycle gas with a vaporous hydrogen-containing stream of dialkyl(e.g. dimethyl) cyclohexanedicarboxylate. The direction of flow of thestream of hydrogen-containing gas through the or each hydrogenation zonein which catalyst reactivation is occurring may be the same as, oropposite to, the direction of flow of the vaporous feed stream throughthat zone when it is on line.

BRIEF DESCRIPTION OF THE DRAWINGS

FIGS. 1 to 4 are each a simplified flow diagram of a plant forproduction of 1,4-cyclohexanedimethanol in two hydrogenation reactorsconnected in parallel by hydrogenation of dimethyl1,4-cyclohexanedicarboxylate; and

FIG. 5 is a simplified flow diagram of an experimental apparatus forproduction of 1,4-cyclohexanedimethanol in a single hydrogenation zoneby hydrogenation of dimethyl 1,4-cyclohexanedicarboxylate.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

It will be understood by those skilled in the art that FIGS. 1 to 4 ofthe drawings are diagrammatic and that further items of equipment suchas temperature and pressure sensors, pressure relief valves, controlvalves, level controllers and the like would additionally be required ina commercial plant. The provision of such ancillary items of equipmentforms no part of the present invention and would be in accordance withconventional chemical engineering practice. Moreover it is not intendedthat the scope of the invention should be limited in any way by theprecise methods of heating, vaporising and condensing various processstreams or by the arrangement of heaters, heat exchangers, vaporising orcondensing apparatus provided therefor. Any suitable arrangement ofequipment other than that depicted in FIGS. 1 to 4 which fulfils therequirements of the invention may be used in place of the illustratedequipment in accordance with conventional chemical engineeringtechniques.

Referring to FIG. 1 of the drawings, a technical grade of dimethyl1,4-cyclohexanedicarboxylate is supplied in line 1, in a first phase ofoperation, to a vaporiser nozzle 2 located in an upper part of avaporiser vessel 3 above a bed of packing 4. A stream of hothydrogen-containing gas is supplied to the bottom of vaporiser vessel 3in line 5. A saturated vaporous mixture comprising dimethyl1,4-cyclohexanedicarboxylate is recovered in line 6 from the top ofvaporiser vessel 3. The resulting vaporous mixture is mixed with furtherhot hydrogen-containing gas from line 7 under the control of valve 8.The combined stream which now has a hydrogen:dimethyl1,4-cyclohexanedicarboxylate molar ratio of 400:1 and is at a pressureof about 63 bar and at a temperature of about 230° C., is fed by way ofvalve 9 and line 10 to a hydrogenation reactor 11 which contains a bedof a pelleted heterogeneous hydrogenation catalyst 12, such as reducedcopper chromite or the chromium-free catalyst designated DRD92/89. Thehydrogenation reaction product mixture exits reactor 11 via line 13 andpasses through valve 14 to enter line 15. The hydrogenation reactionproduct mixture in line 15 is cooled in heat interchanger 16 and theresulting partially condensed mixture passes on in line 17 throughcooler 18 in which it is further cooled. The resulting mixture of gasand condensate flows on in line 19 to a gas-liquid separator 20 fromwhich a mixture of methanol and crude 1,4-cyclohexanedimethanol isrecovered in line 21. The uncondensed gaseous mixture in line 22comprises unreacted hydrogen together with inert gases and methanolvapour and is compressed by means of compressor 23 to give a compressedgas stream in line 24.

The compressed recycled gas in line 24 is combined with make-uphydrogen-containing gas from line 25. The combined mixture in line 26 isheated by passage through heat exchanger 16 and flows on in line 27 toheater 28 in which its temperature is raised further to a suitabletemperature for effecting vaporisation of the dimethyl1,4-cyclohexanedicarboxylate feed. The resulting hot gas in line 29 isthen divided into two streams, one being the stream in line 5 and theother being a stream in line 30. This latter stream is heated further inheater 31 to a temperature of about 240° C. and passes on by way of line32, valve 33 and lines 34 and 35 to the bottom end of a secondhydrogenation reactor 36 which, in this first phase of operation, is inreactivation mode. Reactor 36 contains a charge of hydrogenationcatalyst 37. The hot gas exiting the top of reactor 36 in line 7 isadmixed, as already described above, with the saturated vaporous mixturein line 6 to increase the hydrogen:dimethyl 1,4-cyclohexanedicarboxylatemolar ratio therein and to raise its temperature above its dew point,e.g. at least 5° C. to 10° C. above its dew point.

The plant also includes lines 38 and 39 and valves 40 and 41 both ofwhich are closed in this phase of operation. Line 42 indicates a line bymeans of which a stream containing any "heavies" collecting in thebottom of vaporiser vessel 3 can be drawn off. Reference numeral 43indicates a purge gas line through which a purge gas stream can be takenin order to limit the build up of inert gases in the circulating gas.Such inert gases may enter the plant in the make up gas stream in line25.

After a period of operation the activity of the catalyst charge 12 willhave declined to a point at which reactivation is desirable. Althoughthe reasons for catalyst deactivation have not been clarified, it can bepostulated that a possible cause of this loss of catalyst activity isthe formation of traces of involatile polyesters on the catalyst surfacedue to ester exchange reactions between, for example, dimethyl1,4-cyclohexanedicarboxylate, on the one hand, and1,4-cyclohexanedimethanol, or methyl4-hydroxymethylcyclohexanecarboxylate, which can be postulated to be anintermediate product of the hydrogenation reaction, orhydroxymethylcyclohexylmethyl 1,4-cyclohexanedicarboxylate, which is theester interchange product between dimethyl 1,4-cyclohexanedicarboxylateand 1,4-cyclohexanedimethanol, on the other hand. The resulting di- ortrimeric materials can then undergo further reaction with components ofthe vaporous mixture to cause these oligomeric chains to grow.Polyethers and mixed polyethers-polyesters can also be formed.

Such polymeric by-products on the catalyst surface are susceptible tohydrogenation. Hence reactivation of the catalyst by treatment with ahot hydrogen-containing gas is possible. It has further been shown inthe course of experimental work to investigate the hydrogenation ofdimethyl 1,4-cyclohexanedicarboxylate which forms the background to thepresent invention that, for whatever reason, the passage of a hot streamof hydrogen-containing gas over partially deactivated catalyst has abeneficial effect in at least partially restoring the activity of thecatalyst.

Accordingly in a second phase of operation valve 33 is shut and valve 41is opened, while valve 14 is closed and valve 40 is opened. In this wayhydrogenation reactor 36 with its fresh or reactivated catalyst charge37 is brought on line, whilst reactor 11 goes into reactivation mode andits partially deactivated charge of catalyst 12 is reactivated. In thissecond mode of operation the saturated vaporous mixture in line 6 ismixed with hot hydrogen-containing gas from line 10 to form a vaporousfeed mixture which flows in line 7 through reactor 36 and its catalystcharge 37. The resulting reaction mixture passes by way of lines 35 and38 through valve 40 to line 15. The hot hydrogen-containing gas fromline 32 passes through valve 41 to line 39 and then through line 13 tothe bottom of hydrogenation reactor 11.

When the catalyst charge 37 has become deactivated to some extent thevalves 14, 33, 40 and 41 can be readjusted to switch the flows throughhydrogenation reactors 11 and 36 back to those of the first phase ofoperation.

The above described steps can be repeated as often as may be expedient,bringing the reactors 11 and 36 on line in turn until the reactivationprocedure no longer results in the desired increase in catalyst activityor until the plant has to be shut down for maintenance or other reasons,whereupon the catalyst charges 12 and 37 can be discharged and replacedby fresh charges of catalyst or catalyst precursor.

The make-up gas in line 25 can be a mixture of hydrogen, optional minoramounts of components such as CO and CO₂, and inert gases, such asargon, nitrogen, or methane, containing at least about 70 mole % ofhydrogen. Preferably the make-up gas contains at least 90 mole %, andeven more preferably at least 97 mole %, of hydrogen. The make-up gascan be produced in any convenient manner, e.g. by partial oxidation orsteam reforming of natural gas followed by the water gas shift reaction,and CO₂ absorption, followed possibly by methanation of at least some ofany residual traces of carbon oxides. Pressure swing absorption can beused if a high purity hydrogen make-up gas is desired.

At start up of the plant the reactors 11 and 36 are each charged with acharge of a heterogeneous hydrogenation catalyst precursor, such as acopper chromite catalyst precursor. Preferably, however, the reactors 11and 36 are charged with a chromium-free hydrogenation catalyst, such asDRD92/89. The catalyst precursor is then reduced carefully following thecatalyst supplier's instructions. If the process of EP-A-0301853 is usedto reduce a copper chromite precursor, then both beds of catalyst 12 and37 can be reduced simultaneously. In other cases it may be expedient toreduce the beds 12 and 37 separately. After pre-reduction of thecatalyst precursor hot hydrogen-containing gas is circulated through theplant. When the appropriate inlet temperatures to vaporiser vessel 3 andto reactor 11 have been achieved the flow of dimethyl1,4-cyclohexanedicarboxylate in line 1 is commenced to bring the planton line in the first phase of operation.

In FIG. 2 of the drawings the same reference numerals have been used asin FIG. 1 to denote like items of equipment. Whereas thehydrogen-containing gas flows in the plant of FIG. 1 through thecatalyst bed 12 or 37 in the opposite direction during the reactivationmode from the direction of flow of the vaporous feed stream through thesame bed 12 or 37 in the on line mode, in the plant of FIG. 2 thedirection of gas flow during the catalyst reactivation mode and thatduring the on line mode through a particular bed 12 or 37 are the samein each case.

In the plant of FIG. 2 hot gas from the stream in line 32 can be fedeither via valve 51 and line 52 into line 7 and then through reactor 36or via valve 53 and line 54 into line 10 and then through reactor 11. Ifreactor 11 is in on line mode with valve 8 closed and reactor 36 is inreactivation mode, then valve 51 is adjusted so that most Of the gasfrom line 32 flow through valve 51 into reactor 36 and only sufficientgas passes through valve 53 into line 10 to raise the inlet temperatureto reactor 11 above its dew point. To bring reactor 36 into on line modevalve 9 is closed and valve 8 is opened, whereupon valve 51 is closedsomewhat and valve 53 is opened a corresponding amount to cause most ofthe gas from line 32 to flow through reactor 11 while only sufficientgas passes through valve 51 to satisfy dew point requirements.

In the plant of FIG. 2 any volatile potential catalyst deactivatingmaterials released in the reaction mode do not pass through the on linecatalyst charge and can be recovered in the product stream in line 21.Although preliminary indications are that no such catalyst deactivatingmaterials are released when the unsaturated organic compound beinghydrogenated is dimethyl 1,4-cyclohexanedicarboxylate, this may not bethe case when other unsaturated organic compounds are beinghydrogenated.

FIG. 3 illustrates a further design of hydrogenation plant in accordancewith the invention. In this plant, as in the plant of FIG. 2, thedirection of gas flow through each of the catalyst beds 12 and 37 duringits respective reactivation mode is the same as the direction of flow ofvaporous feed mixture during its on line mode. The reference numeralsused in FIG. 3 indicate the same items of equipment that appear in FIG.1 and also in FIG. 2.

In the plant of FIG. 3 the hydrogen-containing gas supplied in line 5for vaporisation of the incoming dimethyl 1,4-cyclohexanedicarboxylatefeed in line 1 is passed first through one of the catalyst beds 12 or 37in its respective reactivation mode. Hence, when reactor 36 is inreactivation mode, most of the hot gas in line 29 is fed through valve51 and line 52 into reactor 36 through catalyst charge 37, and passesout via line 35 and valve 61 to line 5. Valves 8 and 40 are closed andvalve 53 is open only so far as is necessary to permit passage ofsufficient gas into line 10 to satisfy dew point requirements. Meanwhilethe vaporous feed mixture in line 6 passes through valve 9 and via line10 into reactor 11 whose catalyst charge 12 is in on line mode. Theproduct stream in line 13 flows through valve 14 into line 15, valve 62being closed. When it is desired to bring catalyst charge 37 on line andto reactivate catalyst charge 12 in reactor 11, valves 8, 40 and 62 areopened, while formerly open valves 9, 14 and 61 are closed. Valve 53 isopened somewhat and the gas flow through valve 51 is reduced to theextent necessary for dew point considerations. To revert to the formercondition with catalyst charge 37 being reactivated and catalyst charge12 being on line again, the conditions of the various valves are eachreadjusted to its respective former condition. This procedure can berepeated one or more further times.

In the plant of FIG. 4 part of the hot mixture of recycle gas andmake-up gas in line 29 is fed by way of line 5 to vaporiser 3 while theremainder bypasses vaporiser 3 in line 30 and is admixed with theessentially saturated vaporous hydrogen-containing stream of dimethylcyclohexanedicarboxylate in line 6 so as to form the feed mixture inline 63 for supply to the hydrogenation step. The feed stream in line 63can be divided and fed by way of lines 7 and 10 to hydrogenationreactors 36 and 11 which have outlet lines 35 and 13 which can beconnected simultaneously to common exit line 15. Alternatively theentire vaporous stream in line 63 can be fed through one of reactors 36and 11 only, by way of line 7 or line 10 and the product mixturecollected by way of line 35 or line 13 in line 15.

At start up of the plant of FIG. 4 the hydrogenation reactors 36 and 11are each charged with a charge of a hydrogenation catalyst precursorwhich is then reduced carefully following the catalyst supplier'sinstructions. If using the process of EP-A-0301853 to reduce a copperchromite precursor, then both beds of catalyst 37 and 12 can be reducedsimultaneously. In other cases it may be expedient to reduce the beds 37and 12 separately. After pre-reduction of the catalyst precursor hothydrogen-containing gas is circulated through the plant. When theappropriate outlet temperatures to vaporiser vessel 3 and to reactor 36and/or reactor 11 have been achieved the flow of dimethyl1,4-cyclohexanedicarboxylate in line 1 is commenced.

If the procedure of WO-A-91/01961 is to be used, reactor 36 is usedalone in a first phase of operation, with reactor 11 off line in astandby mode and with its catalyst bed 12 bathed with hydrogen. After asuitable period of operation during which the activity of the catalystin bed 37 declines somewhat, the feed is switched to reactor 11 whilereactor 36 is placed on standby. After a second phase of operation ofthe plant during which the activity of the hydrogenation catalyst in bed12 declines, reactor 36 is brought on line again so that both reactors36 and 11 are used simultaneously in parallel in a third phase ofoperation. In this way the throughput of the plant can be maintained,whilst the corresponding liquid hourly space velocity in each reactor 36or 11 is about 50% of that prevailing during the respective first andsecond phases of operation of the plant. After this third period ofoperation the activity of the catalyst will have declined to a point atwhich it is desirable to shut down the plant for maintenance and toinstall a new catalyst charge.

The invention is further described with reference to the followingExamples. The compositions of catalysts A to D used in the Examples arelisted in Table I. The oxygen content of the catalyst has been excludedfrom the analysis in each case.

                                      TABLE I                                     __________________________________________________________________________                                 Surface   Pore                                           Composition wt %     area Density                                                                            volume                                 Catalyst                                                                              Cu Cr Zn  Mn Ba  Al  m.sup.2 /g                                                                         g/cm.sup.3                                                                         mm.sup.3 /g                            __________________________________________________________________________    A PG 85/1                                                                             42.4                                                                             31.4                                                                             <0.01                                                                             0.02                                                                             0.05                                                                              <0.01                                                                             25   1.275                                                                              260                                    B DRD 89/21                                                                           57.6                                                                             19.0                                                                             <0.01                                                                             0.09                                                                             <0.01                                                                             <0.01                                                                             28   1.420                                                                              200                                    C DRD 92/89                                                                           41.1                                                                             0.26                                                                             <0.01                                                                             6.4                                                                              <0.01                                                                             20.4                                                                              47.1 1.452                                                                              211                                    D DRD 92/92                                                                           34.5                                                                             0.02                                                                             43.1                                                                              2.2                                                                              <0.01                                                                             <0.01                                                                             70   1.423                                                                              210                                    __________________________________________________________________________

EXAMPLE 1

The hydrogenation of a high purity grade of dimethyl1,4-cyclohexanedicarboxylate was investigated using the experimentalapparatus illustrated in FIG. 5.

The composition of the high purity feed was: 36.16 wt % trans-dimethyl1,4-cyclohexanedicarboxylate, 63.26 wt % cis-dimethyl1,4-cyclohexanedicarboxylate, 0.17% methyl hydrogen1,4-cyclohexanedicarboxylate of formula ##STR3## and 0.07 wt % water,with the balance being impurities.

In a commercial plant, hydrogen gas is separated from the hydrogenationproduct and is advantageously recycled through the hydrogenation zone.The hydrogen recycle stream will contain a quantity of methanol vapourproduced by the hydrogenation of dimethyl 1,4-cyclohexanedicarboxylate.Hence, the vaporous mixture supplied to the hydrogenation zone in acommercial plant will generally contain methanol in addition to hydrogenand an unsaturated organic compound. In order that the experimental rigdescribed hereinbelow should accurately predict the likely resultsobtained during commercial operation, the liquid feed supplied to thevaporiser was supplemented by a quantity of liquid methanolcorresponding to the quantity of methanol which would be contained inthe recycle hydrogen stream in a commercial plant. Although hydrogen isrecycled in the experimental rig described hereinbelow, the quantity ofmethanol contained within the recycle hydrogen stream is proportionatelyless than would be contained in a corresponding commercial recyclestream. This difference arises because the recycle gas in theexperimental rig is cooled substantially below the temperature to whichit would be desirably cooled in a commercial plant. More methanol istherefore "knocked out" of the experimental recycle hydrogen stream.This discrepancy between the experimental rig and a commercial plant isnecessitated by the delicacy of the equipment, particularly theanalytical equipment, used in the experimental rig. In this Example andin all succeeding Examples, methanol is added to the experimental liquidfeed in a quantity which is substantially equal to the proportionatequantity of methanol which would be present in the experimental recyclestream if the rig were operated under commercial conditions minus thequantity of methanol actually present in the experimental recyclehydrogen stream. In the Examples, all parameters such as conversionrates and hourly space velocities are calculated on a methanol freebasis.

The experimental apparatus is illustrated in FIG. 5. An approximately 70wt % solution of the high purity grade of dimethyl1,4-cyclohexanedicarboxylate in methanol is fed from reservoir 100 byway of valve 101, line 102 and valve 103 to liquid feed pump 104.Burette 105 provides a buffer supply whilst burette 106 is fitted with aliquid level controller (not shown) that controls valve 101 so as toensure that liquid feed is supplied from reservoir 100 to liquid feedpump 104 at a constant head. The liquid feed is pumped throughnon-return valve 107 and isolation valve 108 into line 109, which can beheated by electrical heating tape 110, before the heated liquid entersthe upper part of an insulated vaporiser vessel 111 above a bed of 6mm×6 mm glass rings 112. A stainless steel demister pad 113 is fitted atthe top end of the vaporiser vessel 111. A stream of hothydrogen-containing gas is supplied to the bottom of vaporiser 111 inline 114. A liquid drain line 115 fitted with a drain valve 116 enableswithdrawal of any unvaporised liquid feed material (e.g. "heavies") fromthe base of the vaporiser vessel 111. The vaporisation of the liquidfeed supplied to the vaporiser vessel 111 is assisted by heating tape117. A saturated vaporous mixture comprising dimethyl1,4-cyclohexanedicarboxylate and hydrogen is recovered in line 118 fromthe top of vaporiser vessel 111. The vaporous mixture is heated byheating tape 119 in order to raise its temperature above the dew pointof the mixture prior to entering the top end of hydrogenation reactor120 which contains a bed of 300 ml (428.1 g) of a pelleted copperchromite hydrogenation catalyst 121. The catalyst was catalyst A ofTable I. Glass rings are packed in reactor 120 above and below thecatalyst bed 121. The vaporous mixture passes downward through catalystbed 121 where conversion of dimethyl 1,4-cyclohexanedicarboxylate to1,4-cyclohexanedimethanol occurs under adiabatic conditions.Adiabaticity is maintained by electrical heating tapes (not shown)within insulating material around reactor 120 under the control ofappropriately positioned thermocouples (not shown). The overall reactionis mildly exothermic with a general increase in catalyst bed temperatureof approximately 1° to 2° C. The hydrogenation product mixture exits thehydrogenation reactor 120 in line 122 and is passed through heatexchanger 123 which simultaneously cools the hydrogenation productmixture and heats a supply of hydrogen-containing gas from line 124.Condensation of the bulk of the 1,4-cyclohexanedimethanol in line 122occurs in heat exchanger 123. The gas in line 124 compriseshydrogen-containing gas from line 125 and, optionally, an inert gas or amixture of inert gases such as nitrogen, argon or methane supplied inline 126. The gas in line 125 comprises make-up hydrogen supplied inline 127 and recycle hydrogen supplied in line 128. Make-up hydrogen inline 127 may be supplied to line 125 in either or both of two streams inlines 129 and 130 via a system of pressure controllers 131 to 136 and amass flow controller 137 from high purity hydrogen cylinders (notshown).

The heated hydrogen-containing gas from heat exchanger 123 passes on inline 114 and is heated further by electrical heating tape 138 for supplyto the vaporiser vessel 111.

The cooled hydrogenation product from heat exchanger 123 passes onthrough line 139 to be cooled further in cooler 140 to a temperaturenear ambient temperature. The liquid/vapour mixture from cooler 140passes on in line 141 to a first knockout pot 142 where liquidhydrogenation product is collected for eventual supply by means of valve143, line 144 and control valve 145 to product line 146. A vaporousmixture comprising hydrogen and uncondensed methanol exits the top ofknockout pot 142 in line 147 and is further cooled to a temperature of10° C. in cooler 148. The further cooled liquid/vapour mixture fromcooler 148 is supplied via line 149 to a second knockout pot 150 whereincondensed methanol is collected for eventual supply through valve 151and line 152 to product line 146. The gas and uncondensed materials fromknockout pot 150 are supplied via line 153 through suction pot 154 intoline 155 and then through valve 156 to gas recycle compressor 157. Gasis recycled through valve 158 lines 128, 125, 124 and 114 to vaporiser111. In order to control the concentration of inert gases, such asnitrogen, in the circulating gas a purge gas stream may be bled from thesystem in line 159 under the control of valve 160.

Reference numeral 161 indicates a bypass valve.

At start up of the apparatus the charge of catalyst was placed inreactor 120 which was then purged with nitrogen. The catalyst charge wasthen reduced according to the teachings of EP-A-0301853.

High purity dimethyl 1,4-cyclohexanedicarboxylate, appropriately dilutedwith methanol, was then pumped to the vaporiser 111 at a rate of 75 ml/hcorresponding to a liquid hourly space velocity of 0.25 h⁻¹. Thegas:dimethyl 1,4-cyclohexanedicarboxylate molar ratio in the vaporousmixture in line 118 was 915:1. The reactor 120 was maintained at atemperature of 220° C. and a pressure of 900 psia (62.05 bar). Thehydrogenation zone was therefore operated under conditions whichprevented the condensation of both dimethyl 1,4-cyclohexanedicarboxylateand the less volatile 1,4-cyclohexanedimethanol product. The temperaturethroughout the hydrogenation zone was above the dew point at theoperating pressure.

The liquid in line 146 was analysed periodically by capillary gaschromatography using a 15 m long, 0.32 mm internal diameter fused silicacolumn coated internally with a 0.25 82 m film of DB wax, a helium flowrate of 2 ml/minute with a gas feed split ratio of 100:1 and a flameionisation detector. The instrument was fitted with a chart recorderhaving a peak integrator and was calibrated using a commerciallyavailable sample of dimethyl 1,4-cyclohexanedicarboxylate of knowncomposition. The exit gas was also sampled and analysed by gaschromatography using the same technique. The identities of the peakswere confirmed by comparison of the retention times observed with thoseof authentic specimens of the materials in question and by massspectroscopy. Included amongst the compounds detected in the reactionmixture were 1,4-cyclohexanedimethanol, dimethyl1,4-cyclohexanedicarboxylate, 4-methoxymethyl cyclohexanemethanol,di-(4-methoxymethylcyclohexylmethyl) ether, and methanol. From theresults obtained it was demonstrated that dimethyl1,4-cyclohexanedicarboxylate can be converted in excess of 99%, with aselectivity to 1,4-cyclohexanedimethanol of approximately 98.5% beingobtained, the balance being minor by-products. After making dueallowance for the methanol present in the feed solution of dimethyl1,4-cyclohexanedicarboxylate from reservoir 100, 2 moles of methanolwere detected for every 1 mole of dimethyl 1,4-cyclohexanedicarboxylateconverted in accordance with the stoichiometry of the hydrogenationreaction. The results are listed in Table II below, together with theresults from the succeeding Examples 2 to 8.

                                      TABLE II                                    __________________________________________________________________________         Pressure                                                                            Inlet                                                                             Gas:DMCD   DMCD  CHDM                                          Example                                                                            psia  Temp.                                                                             molar  LHSV                                                                              conversion                                                                          trans-:cis-                                                                        Selectivity mol %                        No.  (bar) °C.                                                                        ratio  h.sup.-1                                                                          mol % ratio                                                                              CHDM BYPR                                                                              METH                                                                              DETH                        __________________________________________________________________________    1    900 (62.05)                                                                         220 915    0.25                                                                              99.95 3.70 98.35                                                                              1.13                                                                              0.33                                                                              0.19                        2    900 (62.05)                                                                         220 876    0.27                                                                              99.95 3.71 98.47                                                                              1.13                                                                              0.27                                                                              0.13                        3    900 (62.05)                                                                         222 682    0.43                                                                              99.93 3.65 98.66                                                                              0.94                                                                              0.26                                                                              0.14                        4    900 (62.05)                                                                         221 651    0.59                                                                              97.36 2.73 99.09                                                                              0.58                                                                              0.17                                                                              0.16                        5    900 (62.05)                                                                         240 356    0.61                                                                              99.64 3.32 97.99                                                                              1.70                                                                              0.24                                                                              0.07                        6    900 (62.05)                                                                         231 535    0.60                                                                              99.21 3.17 98.84                                                                              0.96                                                                              0.15                                                                              0.05                        7    900 (62.05)                                                                         231 550    0.40                                                                              99.82 3.47 98.39                                                                              1.26                                                                              0.23                                                                              0.12                        8    900 (62.05)                                                                         222 721    0.40                                                                              99.40 3.22 99.26                                                                              0.51                                                                              0.14                                                                              0.09                        __________________________________________________________________________

Notes to Table II:

DMCD=dimethyl 1,4-cyclohexanedicarboxylate

LHSV=liquid hourly space velocity

CHDM=cyclohexanedimethanol

BYPR=miscellaneous by-products

METH=4-methoxymethyl cyclohexanemethanol

DETH=di-4-hydroxymethylcyclohexylmethyl ether

Gas=hydrogen containing gas containing more than 98% hydrogen

EXAMPLES 2 TO 8

Using a similar procedure to that described in Example 1 and the samefeed solution, 7 further runs were carried out using the copper chromitecatalyst (catalyst A in Table I). The runs were designed to investigatethe effect on the hydrogenation reaction of changes in operatingconditions such as temperature, pressure, LHSV and gas:ester mole ratio.In each case the vaporous mixture in contact with the catalyst was aboveits dew point. The results are summarised in Table II.

EXAMPLE 9

The charge of copper chromite catalyst used in Examples 1 to 8 wasreplaced by 150 ml (240.8 g) of DRD 89/21 (catalyst B in Table I), acopper-chromite catalyst with a Cu/Cr weight ratio of 3:1. The catalystwas activated following the procedure described in Example 1 and amethanol solution of high purity dimethyl 1,4-cyclohexanedicarboxylatewas supplied to the vaporiser 111 at a rate of 60 ml/h corresponding toliquid hourly space velocity of 0.40 h⁻¹. The inlet temperature to thehydrogenation zone was 221° C., i.e. above the dew point of the vaporousmixture.

A conversion of 99.52% of dimethyl-1,4-cyclohexanedicarboxylate wasobtained. Detailed results are shown in Table III below, as well as theresults obtained in Examples 10 to 23. The notes to Table II apply alsoto Table III.

                                      TABLE III                                   __________________________________________________________________________         Pressure                                                                            Inlet          DMCD  CHDM                                          Example                                                                            psia  Temp.                                                                             Gas:DMCD                                                                             LHSV                                                                              conversion                                                                          trans-:cis-                                                                        Selectivity mol %                        No.  (bar) °C.                                                                        mol ratio                                                                            h.sup.-1                                                                          mol % ratio                                                                              CHDM BYPR                                                                              METH                                                                              DETH                        __________________________________________________________________________     9   690 (47.57)                                                                         221 543    0.40                                                                              99.52 3.39 98.92                                                                              0.70                                                                              0.21                                                                              0.17                        10   915 (63.09)                                                                         221 773    0.40                                                                              99.66 3.27 99.03                                                                              0.65                                                                              0.20                                                                              0.12                        11   450 (31.03)                                                                         220 346    0.41                                                                              98.23 3.37 98.95                                                                              0.78                                                                              0.17                                                                              0.10                        12   900 (62.05)                                                                         219 684    0.41                                                                              99.70 3.41 99.08                                                                              0.65                                                                              0.19                                                                              0.08                        13   900 (62.05)                                                                         220 697    0.41                                                                              99.57 3.00 98.29                                                                              0.90                                                                              0.72                                                                              0.09                        14   900 (62.05)                                                                         220 695    0.62                                                                              98.21 2.47 99.01                                                                              0.64                                                                              0.25                                                                              0.10                        15   901 (62.12)                                                                         221 659    0.42                                                                              98.71 2.61 96.98                                                                              2.77                                                                              0.18                                                                              0.07                        16   899 (61.98)                                                                         221 725    0.58                                                                              93.99 2.00 96.84                                                                              2.93                                                                              0.14                                                                              0.09                        17   901 (62.12)                                                                         220 720    0.40                                                                              98.30 2.45 96.95                                                                              2.79                                                                              0.16                                                                              0.10                        18   900 (62.05)                                                                         240 362    0.40                                                                              99.82 3.31 95.59                                                                              4.03                                                                              0.31                                                                              0.07                        19   903 (62.26)                                                                         240 364    0.60                                                                              98.75 2.93 95.94                                                                              3.71                                                                              0.23                                                                              0.12                        20   900 (62.05)                                                                         220 691    0.42                                                                              97.53 2.35 96.89                                                                              2.84                                                                              0.14                                                                              0.13                        21   900 (62.05)                                                                         220 720    0.40                                                                              97.85 2.39 96.93                                                                              2.81                                                                              0.14                                                                              0.12                        22   901 (62.12)                                                                         221 530    0.41                                                                              98.20 2.48 96.93                                                                              2.81                                                                              0.14                                                                              0.12                        23   900 (62.05)                                                                         224 446    0.40                                                                              98.92 2.76 96.80                                                                              2.84                                                                              0.16                                                                              0.20                        __________________________________________________________________________

EXAMPLES 10 TO 14

The effect on dimethyl 1,4-cyclohexanedicarboxylate hydrogenation ofaltering the operating conditions described in Example 9 wasinvestigated in 5 further experiments. Detailed results are shown inTable III. In each of Examples 10 to 14 the vaporous mixture in contactwith the catalyst was above its dew point.

EXAMPLES 15 to 23

The high purity dimethyl 1,4-cyclohexanedicarboxylate supplied to thehydrogenation zone in Examples 1 to 14 was replaced with a technicalgrade feed. The composition of the technical grade feed was: 33.95 wt %trans-dimethyl 1,4-cyclohexanedicarboxylate, 61.60 wt % cis-dimethyl1,4-cyclohexanedicarboxylate, 1.59 wt % methyl hydrogen1,4-cyclohexanedicarboxylate, 0.07 wt % water and 2.79 wt % of highboiling impurities including di-4-hydroxymethylcyclohexyl methyl ether.The feed was supplemented with methanol as described in Example 1.Detailed results are shown in Table III. In each of these Examples thevaporous mixture in contact with the catalyst was above its dew point.

EXAMPLE 24

The charge of copper chromite catalyst used in Examples 9 to 23 wasreplaced by 300 ml of DRD 92/89 (catalyst C in Table I), a non-chromiumcatalyst containing copper, manganese and alumina. The catalyst wasactivated by a procedure analogous to that described in Example 1 andhigh purity dimethyl 1,4-cyclohexanedicarboxylate was supplied as amethanol solution to the vaporiser 111 at a rate of 123 ml/hcorresponding to a liquid hourly space velocity of 0.42 h⁻¹ Thegas:ester molar ratio of the vaporous mixture reaching the hydrogenationzone was 703:1 and the hydrogenation zone was maintained at 900 psia(62.05) bar with an inlet temperature of 220° C., i.e. 10° C. above thedew point of the vaporous feed mixture at this pressure.

A dimethyl 1,4-cyclohexanedicarboxylate conversion of 99.78% wasobtained. Detailed results are shown in Table IV below. The notes toTable II apply also to Table IV.

EXAMPLE 25

The effect on hydrogenation of high purity dimethyl1,4-cyclohexanedicarboxylate of altering the operating conditions,specifically the LHSV described in Example 24, was investigated in afurther experiment. Detailed results are shown in Table IV. Again thedimethyl 1,4-cyclohexanedicarboxylate was supplied to the vaporiser as asolution in methanol. The vaporous mixture in contact with the catalystwas about 10° C. above its dew point at the operating pressure.

EXAMPLES 26 TO 32

The high purity dimethyl 1,4-cyclohexanedicarboxylate feed of Examples24 and 25 was replaced with a technical grade feed whose composition hasalready been described in Examples 15 to 23. The effect on hydrogenationof this technical grade feed of altering the operating conditionsdescribed in Example 24 was investigated in 7 further experiments.Detailed results are shown in Table IV; the notes to Table II apply alsoto Table IV. In each of Examples 26 to 32 the vaporous mixture incontact with the catalyst was above its dew point.

                                      TABLE IV                                    __________________________________________________________________________         Pressure                                                                            Inlet          DMCD  CHDM                                          Example                                                                            psia  Temp.                                                                             Gas:DMCD                                                                             LHSV                                                                              conversion                                                                          trans-:cis-                                                                        Selectivity mol %                        No.  (bar) °C.                                                                        mol ratio                                                                            h.sup.-1                                                                          mol % ratio                                                                              CHDM BYPR                                                                              METH                                                                              DETH                        __________________________________________________________________________    24   900 (62.04)                                                                         220 703    0.42                                                                              99.78 3.47 98.18                                                                              0.52                                                                              0.11                                                                              0.19                        25   900 (62.04)                                                                         220 691    0.60                                                                              99.19 3.02 99.29                                                                              0.44                                                                              0.10                                                                              0.17                        26   900 (62.04)                                                                         242 363    0.99                                                                              99.55 3.29 98.23                                                                              1.54                                                                              0.14                                                                              0.09                        27   900 (62.04)                                                                         242 378    1.20                                                                              99.20 3.20 98.17                                                                              1.48                                                                              0.13                                                                              0.22                        28   900 (62.04)                                                                         219 684    0.41                                                                              99.75 3.41 97.13                                                                              2.67                                                                              0.08                                                                              0.12                        29   903 (62.24)                                                                         218 690    0.41                                                                              99.73 3.37 97.14                                                                              2.68                                                                              0.08                                                                              0.10                        30   900 (62.04)                                                                         220 684    0.62                                                                              97.61 2.61 97.53                                                                              2.27                                                                              0.06                                                                              0.14                        31   906 (62.52)                                                                         220 731    0.49                                                                              98.10 2.66 97.65                                                                              2.26                                                                              0.09                                                                              0.00                        32   909 (62.66)                                                                         220 495    0.50                                                                              99.57 3.12 96.92                                                                              2.94                                                                              0.07                                                                              0.07                        __________________________________________________________________________

EXAMPLES 33 TO 35

In an experiment to determine the vapour phase equilibrium trans-:cis-1,4-cyclohexanedimethanol product ratio, an experimental apparatus ofthe type described in Example 1 was packed with 250 ml of catalyst C ofTable I and supplied with a feed of a hydrogenation product of dimethyl1,4-cyclohexanedicarboxylate having the following composition: 48.6 wt %methanol, 1.4 wt % dimethyl 1,4-cyclohexanedicarboxylate, 36.2 wt %trans-1,4-cyclohexanedimethanol, 10.4 wt %cis-1,4-cyclohexanedimethanol, 0.11 wt % 4-methoxymethylcyclohexanemethanol, 0.42 wt % di-(4-hydroxymethylcyclohexylmethyl)ether, 0.5 wt % water and 2.31 wt % of other by-products. Thishydrogenation product thus had a trans:cis- 1,4-cyclohexanedimethanolratio of 3.48:1. The process parameters for Example 33 were as set outin Table V. The product from Example 33 was then passed through thereactor of a second experimental rig of the type described in Example 1which was packed with 250 ml of catalyst C of Table I. The processparameters and product trans-:cid- ratio for this second run are shownin Example 34 in Table VI. The product mixture from Example 34 was fedthrough a third experimental rig of the type described in Example 1,again packed with 250 ml of catalyst C of Table I. The trans-:cis-1,4-cyclohexanedimethanol product ratio resulting from this third run,Example 35, can be seen to have stabilised at around 3.84:1, the vapourphase equilibrium value.

                  TABLE V                                                         ______________________________________                                        Example No.   33        34        35                                          ______________________________________                                        Pressure psia (bar)                                                                         900 (62.05)                                                                             900 (62.05)                                                                             900 (62.05)                                 Inlet temp. (°C.)                                                                    221       221       220                                         H.sub.2 :DMCD ratio                                                                         571       570       570                                         Exit temp. (°C.)                                                                     221       221       220                                         Dew point (°C.)                                                                      213       213       213                                         Residence time (sec)                                                                        4.6       9.3       9.3                                         LHSV (h.sup.-1)                                                                             0.37      0.19      0.19                                        trans-/cis-CHDM                                                                             3.77      3.84      3.84                                        DMCD conversion (%)                                                                         99.99     99.99     100.00                                      ______________________________________                                    

EXAMPLES 36 AND 37

The procedure of Example 1 is repeated using, in place of dimethyl1,4-cyclohexanedicarboxylate, dimethyl 1,2-cyclohexanedicarboxylate anddimethyl 1,3-cyclohexanedicarboxylate respectively. Similar results areobtained.

EXAMPLE 38

The procedure of Example 1 is repeated using in place of catalyst A anequal volume of catalyst D. Similar results are observed.

We claim
 1. A process for the production of cyclohexanedimethanol havinga trans-:cis-isomer ratio greater than about 1:1 by hydrogenation ofdialkyl cyclohexanedicarboxylate having a trans-:cis-isomer ratio lessthan about 1:1 which comprises:(a) providing a hydrogenation zonecontaining a charge of a granular heterogenous ester hydrogenationcatalyst selected from copper-containing catalysts and Group VIIImetal-containing catalysts; (b) supplying to the hydrogenation zone avaporous feed stream containing hydrogen and a hydrogenatable materialcomprising dialkyl cyclohexanedicarboxylate at an inlet temperaturewhich is above the dew point of the mixture; (c) maintaining thehydrogenation zone at a combination of a temperature of from about 150°C. to about 350° C. and a pressure of from about 150 psia (about 10.34bar) to about 2000 psia (about 137.90 bar); (d) passing the vaporousfeed stream through the hydrogenation zone; and (e) recovering from thehydrogenation zone a product stream containing cyclohexanedimethanolhaving a trans-:cis-isomer ratio greater than 1:1;wherein the dialkylcyclohexanedicarboxylate is selected from di-(C₁ to C₄ alkyl)cyclohexanedicarboxylates.
 2. A process according to claim 1, in whichthe dialkyl cyclohexanedicarboxylate is dimethyl1,4-cyclohexanedicarboxylate.
 3. A process according to claim 2, inwhich the dialkyl cyclohexanedicarboxylate is cis-dimethyl1,4-cyclohexanedicarboxylate.
 4. A process according to any claim 3, inwhich the dialkyl cyclohexanedicarboxylate comprises a mixture of thecis- and trans-isomers of dimethyl 1,4-cyclohexanedicarboxylate in whichthe molar ratio of trans-dimethyl cyclohexanedicarboxylate tocis-dimethyl cyclohexanedicarboxylate is in the range of from about0.01:1 to about 1:1.
 5. A process according to claim 4, in which thedialkyl cyclohexanedicarboxylate comprises a mixture of the cis- andtrans-isomers of dimethyl 1,4-cyclohexanedicarboxylate in which themolar ratio of trans-dimethyl cyclohexanedicarboxylate to cis-dimethylcyclohexanedicarboxylate is in the range of from about 0.05:1 to about1:1.
 6. A process according to claim 5, in which the dialkylcyclohexanedicarboxylate comprises a mixture of the cis- andtrans-isomers of dimethyl 1,4-cyclohexanedicarboxylate in which themolar ratio of trans-dimethyl cyclohexanedicarboxylate to cis-dimethylcyclohexanedicarboxylate is in the range of from about 0.5:1 to about0.6:1.
 7. A process according to claim 1, in which thehydrogen-containing gas:dialkyl cyclohexanedicarboxylate mole ratio inthe vaporous mixture is in the range of from about 200:1 to about1000:1.
 8. A process according to claim 1, in which the feed temperatureto the hydrogenation zone is in the range of from about 150° C. to about300° C.
 9. A process according to claim 8, in which the feed temperatureto the hydrogenation zone is in the range of from about 200° C. to about260° C.
 10. A process according to claim 8, in which the feed pressureto the hydrogenation zone is in the range of from about 450 psia (about31.03 bar) to about 1000 psia (about 68.95 bar).
 11. A process accordingto claim 8, in which the catalyst is selected from reduced manganesepromoted copper catalysts, reduced copper chromite catalysts, reducedpromoted copper chromite catalysts, supported palladium/zinc catalysts,chemically mixed copper-titanium catalysts, platinum catalysts andpalladium catalysts.
 12. A process according to claim 11, in which thecatalyst is selected from the reduced forms of copper chromite, promotedcopper chromite, and manganese promoted copper catalysts.
 13. A processaccording to claim 12, in which the catalyst comprises not more thanabout 15% by weight of at least one promoter selected from barium,manganese, and mixtures thereof.
 14. A process according to claim 8, inwhich the vaporous feed stream is supplied at a rate corresponding to aliquid hourly space velocity of the dialkyl cyclohexanedicarboxylate offrom about 0.05 to about 4.0 h⁻¹.
 15. A process according to claim 1, inwhich the dialkyl cyclohexanedicarboxylate is dimethyl1,4-cyclohexanedicarboxylate and in which the product stream of step (e)comprises a 1,4-cyclohexanedimethanol product having atrans.-:cis-isomer ratio of from about 2:1 to about 3.84:1.
 16. Aprocess according to claim 15, in which the product stream of step (e)comprises a 1,4-cyclohexanedimethanol product having a trans-:cis-isomerratio of from about 2.6:1 to about 3.84:1.
 17. A process according toclaim 16, in which the product stream of step (e) comprises a1,4-cyclohexanedimethanol product having a trans-:cis-isomer ratio offrom about 3.1:1 to about 3.84:1.
 18. A process according to claim 1, inwhich the vaporous mixture comprises dimethyl1,3-cyclohexanedicarboxylate.
 19. A process according to claim 1, inwhich the vaporous mixture comprises dimethyl1,2-cyclohexanedicarboxylate.
 20. A vapor phase process for theproduction of cyclohexanedimethanol by hydrogenation of dimethylcyclohexanedicarboxylate which comprises:(a) providing a hydrogenationzone containing a charge of a granular heterogenous ester hydrogenationcatalyst selected from reduced copper chromite catalysts wherein theCu:Cr weight ratio is from about 0.1:1 to about 4:1; reduced, copperchromite catalysts wherein the Cu:Cr weight ratio is from about 0.1:1 toabout 4:1 promoted with from about 0.1% by weight up to 15% by weight ofbarium, manganese or a mixture of barium and manganese; or reducedmanganese promoted copper catalysts wherein the Cu:Mn weight ratio isfrom about 2:1 to about 10:1; (b) supplying to the hydrogenation zone avaporous feed stream containing hydrogen and dimethylcyclohexanedicarboxylate at an inlet temperature which is at or abovethe dew point of the mixture, wherein the dimethylcyclohexanedicarboxylate is supplied to the hydrogenation zone at a ratecorresponding to a liquid hourly space velocity of from 0.05 to about4.0 h⁻¹ ; (c) maintaining the hydrogenation zone at a combination of atemperature of from about 150° C. to about 300° C. and a pressure offrom about 150 psia (about 10.34 bar) to about 2000 psia (about 137.90bar) which maintains both the dimethyl cyclohexanedicarboxylate reactantand the cyclohexanedimethanol product in the vapour phase; and (d)recovering from the hydrogenation zone a vaporous cyclohexanedimethanolproduct wherein the trans-:cis-isomer ratio of the cyclohexanedimethanolproduct is in the range of from about 2:1 to about 3.8:1.
 21. A processaccording to claim 20, in which the pressure within the hydrogenationzone is in the range of about 450 psia (about 31.03 bar) up to about1000 psia (about 68.95 bar).
 22. A process according to claim 21, inwhich the heterogeneous ester hydrogenation catalyst is a manganesepromoted copper catalyst having a Cu:Mn weight ratio of from about 2:1to about 10:1.
 23. A vapor phase process for the production of1,4-cyclohexanedimethanol by hydrogenation of dimethyl1,4-cyclohexanedicarboxylate which comprises:(a) providing ahydrogenation zone containing a charge of a granular heterogenous esterhydrogenation catalyst selected from reduced copper chromite catalystswherein the Cu:Cr weight ratio is from about 0.1:1 to about 4:1;reduced, copper chromite catalysts wherein the Cu:Cr weight ratio isfrom about 0.1:1 to about 4:1 promoted with from about 0.1% by weight upto 15% by weight of barium, manganese or a mixture of barium andmanganese; or reduced manganese promoted copper catalysts wherein theCu:Mn weight ratio is from about 2:1 to about 10:1; (b) supplying to thehydrogenation zone a vaporous feed stream containing hydrogen anddimethyl 1,4-cyclohexanedicarboxylate at an inlet temperature which isat or above the dew point of the mixture, wherein the dimethyl1,4-cyclohexanedicarboxylate is supplied to the hydrogenation zone at arate corresponding to a liquid hourly space velocity of from 0.2 toabout 1.0 h⁻¹ ; (c) maintaining the hydrogenation zone at a combinationof a temperature of from about 210° C. to about 260° C. and a pressureof from about 450 psia (about 31.03 bar) to about 1000 psia (about 68.95bar) which maintains both the dimethyl 1,4-cyclohexanedicarboxylatereactant and the 1,4-cyclohexanedimethanol product in the vapour phase;and (d) recovering from the hydrogenation zone a vaporous1,4-cyclohexanedimethanol product wherein the trans-:cis-isomer ratio ofthe cyclohexanedimethanol product is in the range of from about 2.6:1 toabout 3.8:1.